US 6438430 B1 Abstract A kiln thermal and combustion control. A predictive model is provided of the dynamics of selected aspects of the operation of the plant for modeling the dynamics thereof. The model has at least two discrete models associated therewith that model at least two of the selected aspects, the at least two discrete models having different dynamic responses. An optimizer receives desired values for the selected aspects of the operation of the plant modeled by the model and optimizes the inputs to the model to minimize error between the predicted and desired values. A control input device then applies the optimized input values to the plant after optimization thereof.
Claims(24) 1. A controller for controlling a multi-variable input plant having a plurality of manipulatible variables (MVs) as inputs, and operable to provide an output, comprising:
a predictive model of the dynamics of selected aspects of the operation of the plant for modeling the dynamics thereof, said model having at least two discrete models associated therewith that model at least two of the selected aspects, said at least two discrete models having different dynamic responses and each of said discrete models providing a predicted output of said respective selected aspect;
an optimizer for receiving desired values for the selected aspects of the operation of the plant modeled by said predictive model and said predicted outputs from said predictive model and optimizing the inputs to said predictive model to minimize error between the predicted and desired values; and
a control input device for applying the optimized input values to the plant after optimization thereof.
2. The controller of
3. The controller of
4. The controller of
5. The controller of
6. The controller of
7. The controller of
a plurality of independent sub models, each associated with one of the inputs to the associated model, each of said independent sub models operable to model the associated select aspect as a function of the associated input thereto; and
a summing device for summing the output of said independent sub models for each of said discrete models to provide a single output associated with each of said discrete models that is a function of all of the inputs that are input to each of said associated sub models for each of said discrete models.
8. The controller of
a comparator for comparing for each of said discrete models the associated desired value with the predicted output of said discrete model to generate an error value as the difference therebetween for each of said discrete models; and
an iterative optimization device for incrementally changing the MV inputs over a defined dynamic path for all of said discrete models and inputting them to said model over a plurality of iterative cycles to minimize the error between the desired values and the predicted values.
9. The controller of
10. The controller of
11. The controller of
12. The controller of
13. A method for controlling a multi-variable input plant having a plurality of manipulatible variables (MVs) as inputs, and operable to provide an output, comprising the steps of:
providing a predictive model of the dynamics of selected aspects of the operation of the plant for modeling the dynamics thereof, the model having at least two discrete models associated therewith that model at least two of the selected aspects, the at least two discrete models having different dynamic responses and each of the discrete models providing a predicted output of the respective selected aspect;
receiving in an optimizer desired values for the selected aspects of the operation of the plant modeled by the predictive model and the predicted outputs from the predictive model and optimizing the inputs to the model to minimize error between the predicted and desired values; and
applying the optimized input values to the plant after optimization thereof.
14. The method of
15. The method of
16. The method of
17. The method of
18. The method of
19. The method of
providing a plurality of independent sub models, each associated with one of the inputs to the associated model, each of the independent sub models operable to model the associated select aspect as a function of the associated input thereto; and
summing the output of the independent sub models for each of the discrete models to provide a single output associated with each of the discrete models that is a function of all of the inputs that are input to each of the associated sub models for each of the discrete models.
20. The method of
comparing with a comparator for each of the discrete models the associated desired value with the predicted output of the discrete model to generate an error value as the difference therebetween for each of the discrete models; and
incrementally changing with an iterative optimization device the MV inputs over a defined dynamic path for all of the discrete models and inputting them to the model over a plurality of iterative cycles to minimize the error between the desired values and the predicted values.
21. The method of
22. The method of
23. The method of
24. The method of
Description This application is a Continuation-in-Part of pending U.S. patent application Ser. No. 09/514,733, filed Feb. 28, 2000 entitled “METHOD AND APPARATUS FOR CONTROLLING A NON-LINEAR MILL,” which is a Continuation-in-Part of pending U.S. patent application Ser. No. 09/250,432, filed Feb. 16, 1999 entitled “METHOD AND APPARATUS FOR MODELING DYNAMIC AND STEADY STATE PROCESSES FOR PREDICTION, CONTROL AND OPTIMIZATION,” which is a Continuation of issued U.S. patent application Ser. No. 08/643,464, filed May 6, 1996, now U.S. Pat. No. 5,933,345 issued Aug. 3, 1999, entitled “METHOD AND APPARATUS FOR DYNAMIC AND STEADY STATE MODELING OVER A DESIRED PATH BETWEEN TWO END POINTS.” This invention pertains in general to modeling techniques and, more particularly, to combining steady-state and dynamic models for the purpose of prediction, control and optimization for control of a kiln. Process models that are utilized for prediction, control and optimization can be divided into two general categories, steady-state models and dynamic models. In each case the model is a mathematical construct that characterizes the process, and process measurements are utilized to parameterize or fit the model so that it replicates the behavior of the process. The mathematical model can then be implemented in a simulator for prediction or inverted by an optimization algorithm for control or optimization. Steady-state or static models are utilized in modem process control systems that usually store a great deal of data, this data typically containing steady-state information at many different operating conditions. The steady-state information is utilized to train a non-linear model wherein the process input variables are represented by the vector U that is processed through the model to output the dependent variable Y. The non-linear model is a steady-state phenomenological or empirical model developed utilizing several ordered pairs (U
where P is some parameterization, then the steady-state modeling procedure can be presented as:
where U and Y are vectors containing the U The steady-state model therefore represents the process measurements that are taken when the system is in a “static” mode. These measurements do not account for the perturbations that exist when changing from one steady-state condition to another steady-state condition. This is referred to as the dynamic part of a model. A dynamic model is typically a linear model and is obtained from process measurements which are not steady-state measurements; rather, these are the data obtained when the process is moved from one steady-state condition to another steady-state condition. This procedure is where a process input or manipulated variable u(t) is input to a process with a process output or controlled variable y(t) being output and measured. Again, ordered pairs of measured data (u(I), y(I)) can be utilized to parameterize a phenomenological or empirical model, this time the data coming from non-steady-state operation. The dynamic model is represented as:
where p is some parameterization. Then the dynamic modeling procedure can be represented as:
Where u and y are vectors containing the (u(I),y(I)) ordered pair elements. Given the model p, then the steady-state gain of a dynamic model can be calculated as: Unfortunately, almost always the dynamic gain k does not equal the steady-state gain K, since the steady-state gain is modeled on a much larger set of data, whereas the dynamic gain is defined around a set of operating conditions wherein an existing set of operating conditions are mildly perturbed. This results in a shortage of sufficient non-linear information in the dynamic data set in which non-linear information is contained within the static model. Therefore, the gain of the system may not be adequately modeled for an existing set of steady-state operating conditions. Thus, when considering two independent models, one for the steady-state model and one for the dynamic model, there is a mis-match between the gains of the two models when used for prediction, control and optimization. The reason for this mis-match are that the steady-state model is non-linear and the dynamic model is linear, such that the gain of the steady-state model changes depending on the process operating point, with the gain of the linear model being fixed. Also, the data utilized to parameterize the dynamic model do not represent the complete operating range of the process, i.e., the dynamic data is only valid in a narrow region. Further, the dynamic model represents the acceleration properties of the process (like inertia) whereas the steady-state model represents the tradeoffs that determine the process final resting value (similar to the tradeoff between gravity and drag that determines terminal velocity in free fall). One technique for combining non-linear static models and linear dynamic models is referred to as the Hammerstein model. The Hammerstein model is basically an input-output representation that is decomposed into two coupled parts. This utilizes a set of intermediate variables that are determined by the static models which are then utilized to construct the dynamic model. These two models are not independent and are relatively complex to create. The present invention disclosed and claimed herein comprises, in one aspect thereof, a controller for controlling a multi-variable input plant having a plurality of manipulatible variables (MVs) as inputs, and operable to provide an output. A predictive model is provided of the dynamics of selected aspects of the operation of the plant for modeling the dynamics thereof. The model has at least two discrete models associated therewith that model at least two of the selected aspects, the at least two discrete models having different dynamic responses. An optimizer receives desired values for the selected aspects of the operation of the plant modeled by the model and optimizes the inputs to the model to minimize error between the predicted and desired values. A control input device then applies the optimized input values to the plant after optimization thereof. For a more complete understanding of the present invention and the advantages thereof, reference is now made to the following description taken in conjunction with the accompanying Drawings in which: FIG. 1 illustrates a prior art Hammerstein model; FIG. 2 illustrates a block diagram of the modeling technique of the present invention; FIGS. 3 FIG. 4 illustrates detailed block diagram of the dynamic model utilizing the identification method; FIG. 5 illustrates block diagram of the operation of the model of FIG. 4; FIG. 6 an example of the modeling technique of the present invention utilized in a control environment; FIG. 7 illustrates diagrammatic view of a change between two steady-state values; FIG. 8 illustrates a diagrammatic view of the approximation algorithm for changes in the steady-state value; FIG. 9 illustrates a block diagram of the dynamic model; FIG. 10 illustrates a detail of the control network utilizing the error constraining algorithm of the present invention; FIGS. 11 FIG. 12 illustrates a plot depicting desired and predicted behavior; FIG. 13 illustrates various plots for controlling a system to force the predicted behavior to the desired behavior; FIG. 14 illustrates a plot of the trajectory weighting algorithm of the present invention; FIG. 15 illustrates a plot for the constraining algorithm; FIG. 16 illustrates a plot of the error algorithm as a function of time; FIG. 17 illustrates a flowchart depicting the statistical method for generating the filter and defining the end point for the constraining algorithm of FIG. 15; FIG. 18 illustrates a diagrammatic view of the optimization process; FIG. 18 FIG. 19 illustrates a flowchart for the optimization procedure; FIG. 20 illustrates a diagrammatic view of the input space and the error associated therewith; FIG. 21 illustrates diagrammatic view of the confidence factor in the input space; FIG. 22 illustrates a block diagram of the method for utilizing a combination of a non-linear system and a first principal system; and FIG. 23 illustrates an alternate embodiment of the embodiment of FIG. FIG. 24 illustrates block diagram of a kiln, cooler and preheater; FIG. 25 illustrates a block diagram of a non-linear controlled mill; FIG. 26 illustrates a table for various aspects of the mill associated with the fresh feed and the separator speed; FIG. 27 illustrates historical modeling data for the mill; FIGS. 28 FIG. 29 illustrates non-linear gains for fresh feed responses; FIG. 30 illustrates plots of non-linear gains for separator speed responses; and FIG. 31 illustrates a closed loop non-linear model predictive control, target change for Blaine. FIG. 32 illustrates a detailed block diagram of kiln; FIG. 33 illustrates a detailed block diagram of the MPC controller; FIG. 34 illustrates a block diagram of the separate dynamic models for the kiln model; FIG. 35 illustrates a detailed diagram of each of the separate dynamic models; and FIG. 36 illustrates the optimization function. Referring now to FIG. 1, there is illustrated a diagrammatic view of a Hammerstein model of the prior art. This is comprised of a non-linear static operator model Once the steady-state model is obtained, one can then choose the output vector from the hidden layer in the neural network as the intermediate variable for the Hammerstein model. In order to determine the input for the linear dynamic operator, u(t), it is necessary to scale the output vector h(d) from the non-linear static operator model During the development of a linear dynamic model to represent the linear dynamic operator, in the Hammerstein model, it is important that the steady-state non-linearity remain the same. To achieve this goal, one must train the dynamic model subject to a constraint so that the non-linearity learned by the steady-state model remains unchanged after the training. This results in a dependency of the two models on each other. Referring now to FIG. 2, there is illustrated a block diagram of the modeling method of the present invention, which is referred to as a systematic modeling technique. The general concept of the systematic modeling technique in the present invention results from the observation that, while process gains (steady-state behavior) vary with U's and Y's,( i.e., the gains are non-linear), the process dynamics seemingly vary with time only, (i.e., they can be modeled as locally linear, but time-varied). By utilizing non-linear models for the steady-state behavior and linear models for the dynamic behavior, several practical advantages result. They are as follows: 1. Completely rigorous models can be utilized for the steady-state part. This provides a credible basis for economic optimization. 2. The linear models for the dynamic part can be updated on-line, i.e., the dynamic parameters that are known to be time-varying can be adapted slowly. 3. The gains of the dynamic models and the gains of the steady-state models can be forced to be consistent (k=K). With further reference to FIG. 2, there are provided a static or steady-state model In the static model Systematic Model The linear dynamic model where:
and t is time, a where B is the backward shift operator B(x(t))=x(t−1), t=time, the a The gain of this model can be calculated by setting the value of B equal to a value of “1”. The gain will then be defined by the following equation: The a Since the gain K This makes the dynamic model consistent with its steady-state counterpart. Therefore, each time the steady-state value changes, this corresponds to a gain K Referring now to FIGS. 3 In FIG. 3 Referring now to FIG. 4, there is illustrated a block diagram of a method for determining the parameters a In the technique of FIG. 4, the dynamic model Referring now to FIG. 5, there is illustrated a block diagram of the preferred method. The program is initiated in a block Referring now to FIG. 6, there is illustrated a block diagram of one application of the present invention utilizing a control environment. A plant To facilitate the dynamic control aspect, a dynamic controller During the operation of the system, the dynamic controller Approximate Systematic Modeling For the modeling techniques described thus far, consistency between the steady-state and dynamic models is maintained by rescaling the b Referring now to FIG. 7, there is illustrated a diagram for a change between steady state values. As illustrated, the steady-state model will make a change from a steady-state value at a line Referring now to FIG. 8, there is illustrated a diagrammatic view of the system undergoing numerous changes in steady-state value as represented by a stepped line The approximation is provided by the block 1. Computational Complexity: The approximate systematic model will be used in a Model Predictive Control algorithm, therefore, it is required to have low computational complexity. 2. Localized Accuracy: The steady-state model is accurate in localized regions. These regions represent the steady-state operating regimes of the process. The steady-state model is significantly less accurate outside these localized regions. 3. Final Steady-State: Given a steady-state set point change, an optimization algorithm which uses the steady-state model will be used to compute the steady-state inputs required to achieve the set point. Because of item 2, it is assumed that the initial and final steady-states associated with a set-point change are located in regions accurately modeled by the steady-state model. Given the noted criteria, an approximate systematic model can be constructed by enforcing consistency of the steady-state and dynamic model at the initial and final steady-state associated with a set point change and utilizing a linear approximation at points in between the two steady-states. This approximation guarantees that the model is accurate in regions where the steady-state model is well known and utilizes a linear approximation in regions where the steady-state model is known to be less accurate. In addition, the resulting model has low computational complexity. For purposes of this proof, Equation 13 is modified as follows: This new equation 14 utilizes K The approximate systematic model is based upon utilizing the gains associated with the initial and final steady-state values of a set-point change. The initial steady-state gain is denoted K Substituting this approximation into Equation 13 and replacing u(t−d−1)−u To simplify the expression, define the variable b and g Equation 16 may be written as:
Finally, substituting the scaled b's back into the original difference Equation 7, the following expression for the approximate systematic model is obtained: The linear approximation for gain results in a quadratic difference equation for the output. Given Equation 20, the approximate systematic model is shown to be of low computational complexity. It may be used in a MPC algorithm to efficiently compute the required control moves for a transition from one steady-state to another after a set-point change. Note that this applies to the dynamic gain variations between steady-state transitions and not to the actual path values. Control System Error Constraints Referring now to FIG. 9, there is illustrated a block diagram of the prediction engine for the dynamic controller
With further reference to FIG. 9, the input values u(t) for each (u,y) pair are input to a delay line The a In FIG. 9, the coefficients in the coefficient modification block Referring now to FIG. 10, there is illustrated a block diagram of the dynamic controller and optimizer. The dynamic controller includes a dynamic model The output of model The optimization function is performed by the inverse model Referring now to FIGS. 11 Referring now to FIG. 12, there is illustrated a plot of y where: Du A y y Trajectory Weighting The present system utilizes what is referred to as “trajectory weighting” which encompasses the concept that one does not put a constant degree of importance on the future predicted process behavior matching the desired behavior at every future time set, i.e., at low k-values. One approach could be that one is more tolerant of error in the near term (low k-values) than farther into the future (high k-values). The basis for this logic is that the final desired behavior is more important than the path taken to arrive at the desired behavior, otherwise the path traversed would be a step function. This is illustrated in FIG. 13 wherein three possible predicted behaviors are illustrated, one represented by a curve In Equation 23, the predicted curves Error Constraints Referring now to FIG. 15, there are illustrated constraints that can be placed upon the error. There is illustrated a predicted curve The difference between constraint frustums and trajectory weighting is that constraint frustums are an absolute limit (hard constraint) where any behavior satisfying the limit is just as acceptable as any other behavior that also satisfies the limit. Trajectory weighting is a method where differing behaviors have graduated importance in time. It can be seen that the constraints provided by the technique of FIG. 15 requires that the value y Trajectory weighting can be compared with other methods, there being two methods that will be described herein, the dynamic matrix control (DMC) algorithm and the identification and command (IdCom) algorithm. The DMC algorithm utilizes an optimization to solve the control problem by minimizing the objective function: where B It is noted that the weights A The IdCom algorithm utilizes a different approach. Instead of a constant desired value, a path is defined for the control variables to take from the current value to the desired value. This is illustrated in FIG. This technique is described in Richalet, J., A. Rault, J. L. Testud, and J. Papon; Model Predictive Heuristic Control: Applications to Industrial Processes, Automatica, 14, 413-428 (1978), which is incorporated herein by reference. It should be noted that the requirement of Equation 25 at each time interval is sometimes difficult. In fact, for control variables that behave similarly, this can result in quite erratic independent variable changes due to the control algorithm attempting to endlessly meet the desired path exactly. Control algorithms such as the DMC algorithm that utilize a form of matrix inversion in the control calculation, cannot handle control variable hard constraints directly. They must treat them separately, usually in the form of a steady-state linear program. Because this is done as a steady-state problem, the constraints are time invariant by definition. Moreover, since the constraints are not part of a control calculation, there is no protection against the controller violating the hard constraints in the transient while satisfying them at steady-state. With further reference to FIG. 15, the boundaries at the end of the envelope can be defined as described hereinbelow. One technique described in the prior art, W. Edwards Deming, “Out of the Crisis,” Massachusetts Institute of Technology, Center for Advanced Engineering Study, Cambridge Mass., Fifth Printing, September 1988, pages 327-329, describes various Monte Carlo experiments that set forth the premise that any control actions taken to correct for common process variation actually may have a negative impact, which action may work to increase variability rather than the desired effect of reducing variation of the controlled processes. Given that any process has an inherent accuracy, there should be no basis to make a change based on a difference that lies within the accuracy limits of the system utilized to control it. At present, commercial controllers fail to recognize the fact that changes are undesirable, and continually adjust the process, treating all deviation from target, no matter how small, as a special cause deserving of control actions, i.e., they respond to even minimal changes. Over adjustment of the manipulated variables therefore will result, and increase undesirable process variation. By placing limits on the error with the present filtering algorithms described herein, only controller actions that are proven to be necessary are allowed, and thus, the process can settle into a reduced variation free from unmerited controller disturbances. The following discussion will deal with one technique for doing this, this being based on statistical parameters. Filters can be created that prevent model-based controllers from taking any action in the case where the difference between the controlled variable measurement and the desired target value are not significant. The significance level is defined by the accuracy of the model upon which the controller is statistically based. This accuracy is determined as a function of the standard deviation of the error and a predetermined confidence level. The confidence level is based upon the accuracy of the training. Since most training sets for a neural network-based model will have “holes” therein, this will result in inaccuracies within the mapped space. Since a neural network is an empirical model, it is only as accurate as the training data set. Even though the model may not have been trained upon a given set of inputs, it will extrapolate the output and predict a value given a set of inputs, even though these inputs are mapped across a space that is questionable. In these areas, the confidence level in the predicted output is relatively low. This is described in detail in U.S. patent application Ser. No. 08/025,184, filed Mar. 2, 1993, which is incorporated herein by reference. Referring now to FIG. 17, there is illustrated a flowchart depicting the statistical method for generating the filter and defining the end point
where: e a=actual value p=model predicted value The model accuracy is defined by the following equation:
where: Acc=accuracy in terms of minimal detector error σ The program then flows to a function block
where: e d=desired value m=measured value The program will then flow to a decision block
where: S S k=Tuning factor—minimal detectable change threshold with the following defined: Hq=significance level. Values of (j,k) can be found so that the CUSUM control chart will have significance levels equivalent to Shewhart control charts. The program will then flow to a decision block Referring now to FIG. 18, there is illustrated a block diagram of the overall optimization procedure. In the first step of the procedure, the initial steady-state values {Y Referring now to FIG. 18 Referring now to FIG. 19, there is illustrated a flow chart depicting the optimization algorithm. The program is initiated at a start block Steady State Gain Determination Referring now to FIG. 20, there is illustrated a plot of the input space and the error associated therewith. The input space is comprised of two variables x Once the system is operating outside of the training data regions, i.e., in a low confidence region, the accuracy of the neural net is relatively low. In accordance with one aspect of the preferred embodiment, a first principles model g(x) is utilized to govern steady-state operation. The switching between the neural network model f(x) and the first principle models g(x) is not an abrupt switching but, rather, it is a mixture of the two. The steady-state gain relationship is defined in Equation 7 and is set forth in a more simple manner as follows: A new output function Y(u) is defined to take into account the confidence factor α(u) as follows:
where: α(u)=confidence in model f(u) α(u) in the range of 0→1 α(u)ε{0,1} This will give rise to the relationship: In calculating the steady-state gain in accordance with this Equation utilizing the output relationship Y(u), the following will result: Referring now to FIG. 22, there is illustrated a block diagram of the embodiment for realizing the switching between the neural network model and the first principles model. A neural network block Referring now to FIG. 23, there is illustrated an alternate embodiment which utilizes discreet switching. The output of the first principles block The switch Non-linear Mill Control Overall, model predictive control (MPC) has been the standard supervisory control tool for such processes as are required in the cement industry. In the cement industry, particulate is fabricated with a kiln/cooler to generate raw material and then to grind this material with a mill. The overall kiln/cooler application, in the present embodiment, utilizes a model of the process rather than a model of the operator. This model will provide continuous regulation and disturbance rejection which will allow the application to recover from major upsets, such as coating drop three times faster than typical operator intervention. In general, mills demonstrate a severe non-linear behavior. This can present a problem in various aspects due to the fact that the gains at different locations within the input space can change. The cement kilns and coolers present a very difficult problem, in that the associated processes, both chemical and physical, are in theory simple, but in practice complex. This is especially so when commercial issues such as quality and costs of production are considered. The manufacturing of cement, and its primary ingredient, clinker, has a number of conflicting control objectives, which are to maximize production, minimize costs, and maximize efficiency, while at the same time maintaining minimum quality specifications. All of this optimization must take place within various environmental, thermodynamic and mechanical constraints. A primary technique of control for clinker has been the operator. As rotary cement kilns and automation technology evolve, various automation solutions have been developed for the cement industry. These solutions have been successful to a greater or lessor extent. In the present application, the process is modeled, rather than the operator, and model predictive control is utilized. Moves are made every control cycle to the process based on continuous feedback of key measurements. This gives rise to a continuous MPC action, as opposed to the intermittent, albeit frequent moves made by the typical expert system. In addition, as will be described hereinbelow, the approach described utilizes full multivariable control (MVC) techniques, which take into account all coupled interactions in the kiln/cooler process. The cement mill is utilized to manufacture the various grades of cement after processing of the raw material, which are defined by their chemical composition and fineness (particle size distribution). The control objectives are thus to maximize production at minimum cost, i.e., low energy consumption for the various product grades, chemical compositions and specified fineness. In general, the mill utilizes a closed circuit where separators in the feed-back are utilized to classify the mill output into oversized and undersized product streams. The oversized stream, which does not conform to specification required for correct cement strength, is fed back into the mill for further grinding and size reduction. Depending upon the type of mill, controls include fresh feed, recirculating-load, as well as separator speed, all of which are used by the operator to control fineness, energy consumption and throughput. In general, the mill grinding equations take the form of:
where: P=particle size F=feed rate k It has generally been stated in the literature that grinding model equations are non-linear and hence, direct application of linear control theory is not possible. The primary reason for this is that the operation of the plant is only non-linear in very small regions of the input space. Once the process has traversed, i.e., “stepped,” from one portion of the input space to another portion thereof, the overall model changes, i.e., it is non-linear. This lends itself to non-linear modeling techniques. However, most control systems are linear in nature, especially those that model the dynamics of the plant. Referring now to FIG. 24, there is illustrated a diagrammatic view of the kiln/cooler configuration and the selected instrumentation utilized for optimal MPC control. This kiln/cooler consists of a five-stage suspension pre-heater kiln, with back-end firing (approximately fifteen percent of total firing). The cooler is a grate type with a conversion upgrade on the first section. It has on-line analyzers for NOx, O The system is controlled with an MPC controller (not shown) that consists of the MPC control engine as described hereinabove, as well as a real-time expert system that performs a variety of pre and post processing of control signals as well as various other functions such as MPC engine control, noise filtering, bias compensation and real-time trending. This system will also perform set point tracking for bumperless transfer, and adaptive target selection. This allows for the controller tuning parameters to be changed according to various business and/or process strategies. The MPC control solution as discussed hereinbelow, is designed to achieve two primary control goals, the first being optimal operation and stabilization of the kiln/cooler, and the second being clinker quality control. This control must also achieve an optimal balance between quality and production, which implies tying these processes together and then reducing costs in the presence of constraints. This multivariate problem and optimization together with extreme variability (chaotic) plant behavior motivates the use of MPC where the emphasis switches from modeling operators to modeling the basic processes. In this context the disturbances and the problems that could crop up in operations are compensated for automatically due the inherent structure of the MPC. This results in a powerful concept because unlike fuzzy logic, much less effort is required to define how to react to a problem, and the problem is reduced to mathematics where error reduction is the objective. This is in contract to rules and fuzzy logic where given the large number of permutations of plant states, problems and the method to handle them, the definition, management and robustness of rules becomes a problem itself. For example the definition of “hot” for say the hood temperature itself becomes a problem because it depends on the plant state and could have multiple definitions, and given a time trajectory may even fall out of the definition of hot. In this regard a model of the process is required and hence MPC. Kiln process are nevertheless quite complex in their own right, and with MPC the model will inherently be only an approximation. However, the inherent robustness of MPC compensates quite well for model mismatches, and such mismatches can be identified and corrected quite easily. Disturbance Rejection and Dynamics Clinker quality is determined primarily by the characteristics of the kiln feed, as well as kiln thermodynamic conditions. Thus, stable control of the chemical composition of the kiln feed e.g., LSF, and the thermodynamic profile of the plant are necessary requirements. LSF deviations in particular can be a major source of disturbances as LSF effects the “burnability” and hence the thermodynamics. In such cases it is usually only an hour after receipt of lab results that the operator will know why temperatures dropped and, for example, free-limes increased. It is also highly likely that compensation was not optimal because the results of alterative solutions were simply not known, and conventional kiln operations dictate corrective action that errs on the side of comfort. Because of forward predictions into the future, as will be described hereinbelow MPC is particularly good at this type of disturbance rejection, as corrective actions can be taken before hand and the solution will exploit the dynamics of the process to bring variables back on target as fast as possible. For example, the MPC controller may not have increased coal to compensate for dropping temperatures, but rather maximized the use of secondary air from the cooler, a possibly faster and more economic reaction to increasing fuel. In addition, coal set point moves introduce more disturbances into the kiln, and the forward prediction in MPC would include this because of the MPC knowledge of the dynamics. Efficiency and Secondary Air The thermodynamic processes for clinker production are the preheating and calcining of the raw meal, followed by the endothermic, and exothermic sintering processes where the various complexes of clinker such as calcium-silicates (C The reduction of energy consumption is a dominant objective in these processes, and the preheater and cooler are vital to achieving this objective. The preheater is used to exploit as much of the heat from the kiln before being wasted to atmosphere, (calcining). Similarly, maximum effort is used to recover as much of the energy released from the cooling of the product and to re-introduce it back into the kiln, as “secondary” air. With grate coolers, the leveraging of secondary air becomes a sub objective to the maximization of this energy recovery, the control of which is primarily through the grate speed and grate air flow. However the dynamics in the kiln greatly affect the cooler and vice versa. The ID-fan also plays an important role, but the ID-fan also has another major role, which is the control of the combustion process, which is indicated by the O It has been recognized for the purposes of the present disclosure that counter current flow of air and material in the preheater, kiln and cooler form a number of complex feed-forward and feedback thermodynamic process loops, which are difficult to understand and hence control. Compounding this issue is the number of dead-times in these loops. For example, the residence time in the kiln is about 45 minutes, while that of the cooler is about 15 minutes. Conversely some of the process response times are of the order of seconds. This large range of varying time constants and dead time makes it very difficult for the operator to take into account and remember the results of certain moves that will take effect far into the future. An example of this complexity is a scenario where the operator increases kiln feed in order to increase production. The greater mass of cooler material, which begins to flow into the system, will first cause preheater temperatures to drop almost immediately, and hood temperatures to rise and then drop some time later. This scenario is extremely simplistic and is made on the assumption that all other control variables remain constant. In reality the experienced operator or any other advanced control system will try to compensate for this by, for example, increasing the ID-fan and adjusting other control variables such as increasing coal to compensate for the greater energy requirement. In addition, the kiln speed will be increased in order to maintain a proper bed depth, all of which create major fluctuations to the material flow in the cooler, and hence the amount of air and heat being introduced into the kiln as well as secondary air, which itself becomes a major feed-back disturbance the kiln heat balance. These process fluctuations do not only occur during the period of time, the duration of which is related to the dead times and time constants associated with these processes, but are observed to repeat a number of times afterwards, due to feedback. These effects will with time be lost either in the process noise or in the fluctuations of newer disturbances. Also depending on the operational state of the kiln these disturbances could cause the whole heat balance to become unstable, thus causing extremely high or low temperatures, and in general operator over-compensation. It has also been noticed that the kiln/cooler is a set of complex interacting processes and taking the correct and optimal control actions becomes very difficult. The MPC control technique of the present disclosure is that, given a robust multivariable controller, as well as models of the fundamental process responses, the kiln can be controlled more rigorously, and most disturbances measured and unmeasured taken into cognizance fairly far into the future. This helps to stabilize the kiln, minimize operator over-compensation, as well as ensuring fast response to disturbances, thus reducing reject material and lowering fuel consumption. The MPC is defined in two primary phases the first being the modeling phase in which the models of the kiln processes are developed. The second phase is the deployment phase, where the models are commissioned and refined to a point where satisfactory control can be obtained. Central to commissioning is the tuning where the controller is tweaked to provide the desired control and optimization. For example this could be: maximum production at the expense of quality, or optimal quality at the expense of efficiency. The MPC models are developed from the analysis of test and process data, together with knowledge from the plant operators and other domain experts. The result is a matrix of time responses, where each response reflects the dynamic interaction of a controlled variable to a manipulated variable. The tuning involves the selection of targets (set points), weighting factors and various constraints for each variable. This determines how the controller will solve the control problem at any given time. The control of the kiln and its optimization within the above set of constraints is solved every control cycle. The solution chosen in a particular control cycle may not seem to be necessarily optimal at that given time, but will be optimal within the solution space which has temporal as well as spatial dimensions. Thus the control solution is a series of trajectories into the future, where the whole solution is optimized with time. The very nature of optimal control in real time does not allow for a guarantee of a global optimal solution. However the calculation of an optimal solution within a finite amount of time is itself a class of optimization. Some of the tuning parameters, which can be changed during operations, include: 1) Targets. Targets can be set for both controlled and manipulated variables, and the MPC controller will try and force all variables to their desired targets. In particular setting a target for a manipulated variable such as coal allows for optimization, and in particular efficiency, because the controller will continually seek a lower coal flow while maintaining production and quality. For some variables such as O 2) Priorities. Setting relative priorities between manipulated variables and controlled variables allows the controller to prioritize which are more important problems to solve, and what type of solution to apply. Underspecified multivariable control (more manipulated variables than controlled variables, as is the case in this application) implies that for every problem there will be more than one solution, but within constraints one solution will generally be more optimal than others. For example, too high a hood temperature can be controlled by, (a) reducing fuel, (b) increasing the grate speed, or (c) increasing cooler airflow, or a combination of the above. 3) Hard Constraints. Setting upper and lower hard constraints for each process variable, for example, minimum and maximum grate speed. These values which are usually defined by the mechanical and operational limitations of the grate. Maintaining these constraints is obviously realizable with controlled variables such as ID-fan speed, but is more difficult to achieve with, for example, hood temperature. However when hood temperature exceeds a upper hard constraint of say 1200° C., the controller will switch priority to this temperature excursion, and all other control problems will “take a back seat” to the solution required to bring this temperature back into the allowable operating zone. 4) Soft upper and lower constraints. If any process variable penetrates into the soft constraint area, penalties will be incurred that will begin to prioritize the solution of this problem. Continuous penetration into this area will cause increasing prioritization of this problem, thus in effect creating an adaptive prioritization, which changes with the plant state. 5) Maximum rate of change constraints. These parameters are only applicable to the manipulated variables, and generally reflect a mechanical of physical limitation of the equipment used, for example maximum coal feed rate. From a clinker production point of view the functions of the MPC application can be viewed as follows: 1) Kiln Combustion Control where manipulated variables such as ID-fan speed and fuel-flow are manipulated to control primarily O 2) Kiln Thermal “Hinge Point” Control adjusts total coal, cooler grate speed, and cooler fans to control the hood temperature. The hood temperature is conceptualized as the “hinge” point on which the kiln temperature profile hangs. The controller is tuned to constantly minimize cooler grate speed and cooler fans, so that heat recovery from the cooler is maximized, while minimizing moves to coal. 3) Kiln Thermal “Swing Arm” Control adjusts percent coal to the kiln backend, in order to control clinker free lime based on hourly lab feedback. This control function is about three times slower than the hinge point control, which maintains hood temperature at a fixed target. The “swing arm effect” raises or lowers the back end temperature with a constant firing end temperature to compensate for changes in free lime. This is in effect part of the quality control. Kiln combustion control, kiln thermal hinge point control, and kiln thermal swing arm control are implemented in a single MPC controller. Kiln speed is included as a disturbance variable, as the production philosophy, in one embodiment, calls for setting a production rate to meet various commercial obligations. This means that any changes to kiln speed and hence production rate by the operator will be taken into account in the MPC predictions, but the MPC controller will not be able to move kiln speed. The control system allows for customization of the interface between the plant and the MPC special control functions, the special control functions implemented including: 1) Total Coal Control allows the operator to enter a total coal or fuel flow set point. The control system “wrapper” splits the move to the front and back individual coal flow controllers while maintaining the percent of coal to the back constant. The purpose of this control function is to allow heating and cooling of the kiln while maintaining a constant energy profile from the preheaters through to the firing end of the kiln. This provides a solid basis for the temperature “hinge point” advanced control function previously described. 2) Percent Coal to the Back Control allows the operator to enter a percent coal to the back target and implements the moves to the front and back coal flow controllers to enforce it. The purpose of this control is to allow the back end temperature to be swung up or down by the thermal “swing arm” advanced control function. 3) Feed-to-Speed Ratio Control adjusts raw meal feed to the kiln to maintain a constant ratio to kiln speed. The purpose of this controller is to maintain a constant bed depth in the kiln, which is important for long-term stabilization. 4) Cooler Fans Control is a move splitter that relates a single generic cooler air fans set point to actual set points required by n cooler air fans. The expert system wrapper through intelligent diagnostics or by operator selection can determine which of the air cooler fans will be placed under direct control of the MPC controller, thus allowing for full control irrespective of the (for example) maintenance being undertaken on any fans. 5) Gas analyzer selection. The control system automatically scans the health of the gas analyzers, and will switch to the alternative analyzer should the signals become “unhealthy”. In addition the control system is used to intelligently extract the fundamental control signals from the O Referring now to FIG. 25, there is illustrated a block diagram of the non-linear mill and the basic instrumentation utilized for advanced control therein. The particle size overall is measured as “Blaine” in cm There are provided various sensors for the operation of the mill. The separator speed is controlled by an input Overall, the mill-separator-return system is referred to as a “mill circuit.” The main control variable for a mill circuit is product particle size, the output, and fresh feed is manipulated to control this variable. A secondary control variable is return and separator speed is manipulated to control this variable. There are also provided various constants as inputs and constraints for the control operation. This controller The response of particle size to a move in fresh feed is known to be slow (one-two hours) and is dominated by dead time. Where a dead time to time constant ratio exceeding 0.5 is known to be difficult to control without model predictive control techniques, documents ratios for the response of particle size to a move in fresh feed includes 0.9 and 1.3. In the case of a closed-circuit mill, a move to fresh feed effects not only the product particle size, but also the return flow. Also, a move to separator speed effects not only the return flow, but also the particle size. This is a fully interactive multi-variable control problem. The controller adjusts fresh feed and separator speed to control Blaine and return. It also includes motor and sonic ear as outputs, and the sonic ear is currently used as a constraint variable. That means when the sonic ear decibel reading is too high then fresh feed is decreased. In this way the controller maximizes feed to the sonic ear (choking) constraint. Referring now to FIG. 26, there is illustrated a dynamic model matrix for the fresh feed and the separator speed for the measured variables of the Blaine Return Ear and motor. It can be seen that each of these outputs has a minimum and maximum gain associated therewith dead-time delay and various time constants. Referring now to FIG. 27, there is illustrated a plot of log sheet data utilized to generate a gain model for the controller In general, the operation described hereinabove utilizes a non-linear controller which provides a model of the dynamics of the plants in a particular region. The only difference in the non-linear model between one region of the input space to a second region of the input space is that associated with the dynamic gain “k.” This dynamic gain varies as the input space is traversed, i.e., the model is only valid over a small region of the input space for a given dynamic gain. In order to compensate for this dynamic gain of a dynamic linear model, i.e., the controller, a non-linear steady state model of the overall process is utilized to calculate a steady-state gain “K” which is then utilized to modify the dynamic gain “k.” This was described in detail hereinabove. In order to utilize this model, it is necessary to first model the non-linear operation of the process, i.e., determining a non-linear steady state model, and then also determine the various dynamics of the system through “step testing.” The historical data provided by the log sheets of FIG. 27 provide this information which can then be utilized to train the non-linear steady state model. Referring now to FIGS. 28 In operation of the controller Referring now to FIG. 29, there is illustrated a plot of non-linear gains for fresh feed responses for both manipulatible variables and control variables. The manipulatible variable for the fresh feed is stepped from a value of 120.0 to approximately 70.0. It can be seen that the corresponding Blaine output and the return output, the controlled variables, also steps accordingly. However, it can be seen that a step size of 10.0 in the fresh feed input does not result in identical step in either the Blaine or the return over the subsequent steps. Therefore, it can be seen that this is clearly a non-linear system with varying gains. Referring now to FIG. 30, there is illustrated four plots of the non-linear gains for separator speed responses, wherein the separator speed as a manipulatible variable is stepped from a value of approximately 45.0 to 85.0, in steps of approximately 10.0. It can be seen that the return controlled variable makes an initial change with an initial step that is fairly large compared to any change at the end. The Blaine, by comparison, appears to change an identical amount each time. However, it can be seen that the response of the return with respect to changes in the separator speed will result in a very non-linear system. Referring now to FIG. 31, there is illustrated an overall output for the closed loop control operation wherein a target change for the Blaine is provided. In this system, it can be seen that the fresh feed is stepped at a time Referring now to FIG. 32, there is illustrated a more detailed schematic of the kiln The cylinder As the cylinder At the back end The raw meal (powder) is provided on a feed mechanism During control, as described hereinabove, there is provided a level of control that can control the total coal that is fed to the system, maintain the ratio of the back coal and the front coal to each other, maintain the O Referring now to FIG. 33, there is illustrated a block diagram of the overall MPC as it applies to the overall kiln. In the illustration of FIG. 33, there is provided a general a block A controller In addition to the sensed values s(t), there is also provided a separate input that is a function of the free lime measurement. This free lime measurement is a lab measurement wherein the actual output—the clinker, will be evaluated to determined the percent free lime therein. This percent free lime measurement is input to a free lime model The output of the kiln model Referring now to FIG. 34, there is illustrated a simplified diagrammatic view of the kiln model Referring now to FIG. 35, there is illustrated a more detailed block diagram of a single one of the models, the example illustrating that of the model Each of the outputs of each of the models
Referring now to FIG. 36, there is illustrated a block diagram for the optimizing operation. Each of the outputs of each of the models Although the preferred embodiment has been described in detail, it should be understood that various changes, substitutions and alterations can be made therein without departing from the spirit and scope of the invention as defined by the appended claims. Patent Citations
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